Process and system for reducing the olefin content of a fischer-tropsch product stream

ABSTRACT

A method is provided for converting synthesis gas to liquid hydrocarbon mixtures useful as distillate fuel and/or lube base oil. The synthesis gas is contacted with a synthesis gas conversion catalyst comprising a Fischer-Tropsch synthesis component in an upstream catalyst bed thereby producing an intermediate hydrocarbon mixture containing olefins and C 21+  normal paraffins. The intermediate hydrocarbon mixture is subsequently contacted with a hydroisomerization catalyst and an olefin saturation catalyst, thereby resulting in a product containing no greater than about 25 wt % olefins and containing no greater than about 5 wt % C 21+  normal paraffins. The hydroisomerization and olefin saturation catalysts may be in separate beds or mixed in a single bed downstream of the synthesis gas conversion catalyst.

FIELD

The present invention relates to a process for converting synthesis gas to liquid hydrocarbon mixtures useful as distillate fuel and/or lube base oil, the process including contacting the synthesis gas with a catalyst thereby producing a liquid containing a paraffin component and an olefin component, and saturating the olefin component by contacting the liquid with a hydrogenation catalyst.

BACKGROUND

The majority of combustible liquid fuel used in the world today is derived from crude oil. However, there are several limitations to using crude oil as a fuel source. For example, crude oil is in limited supply.

Alternative sources for developing combustible liquid fuel are desirable. An abundant resource is natural gas. The conversion of natural gas to combustible liquid fuel typically involves a first step of converting the natural gas, which is mostly methane, to synthesis gas, or syngas, which is a mixture of carbon monoxide and hydrogen. Fischer-Tropsch synthesis is a known means for converting syngas to higher molecular weight hydrocarbon products. Fischer-Tropsch diesel has a very high cetane number and is effective in blends with conventional diesel to reduce NO_(x) and particulates from diesel engines, allowing them to meet stricter emissions standards.

Fischer-Tropsch synthesis is often performed under conditions which produce a large quantity of C₂₁₊ wax, also referred to herein as “Fischer-Tropsch wax,” which must be hydroprocessed to provide distillate fuels. Often, the wax is hydrocracked to reduce the chain length, and then hydrotreated to reduce oxygenates and olefins to paraffins. Hydrocracking tends to reduce the chain length of all of the hydrocarbons in the feed. When the feed includes hydrocarbons that are already in a desired range, for example, the distillate fuel range, hydrocracking of these hydrocarbons is undesirable.

As disclosed in co-pending U.S. patent application Ser. No. 12/343,534, incorporated in its entirety by reference, a hybrid Fischer-Tropsch catalyst, also referred to herein as a hybrid synthesis gas conversion catalyst, is described which is capable of converting synthesis gas to a hydrocarbon mixture free of solid wax. One advantage of a process employing this catalyst is that the absence of a solid wax phase eliminates the need for separating, and hydrotreating and/or hydrocracking a waxy product in a separate reactor. As such the hydrocarbon product resulting from this improved process can, in theory, be blended with crude oil.

In practice, however, Fischer-Tropsch synthesis produces a large percentage of olefinic hydrocarbons. An olefinic hydrocarbon is defined as a hydrocarbon in which one or more double bonds exist within the molecule. Olefinic, or unsaturated, hydrocarbons have the potential to be disruptive to refining processes, creating problems including crude heater and preheat train fouling, storage instability and gum deposits. Furthermore, the hydrogenation of olefins, apart from diene saturation, is not practiced in crude oil refining. For this reason, synthetic hydrocarbon mixtures must be treated so as to substantially remove unsaturated hydrocarbons before being blended into crude oil.

It would be desirable to have a means for converting synthesis gas to a hydrocarbon mixture free of solid wax with a low percentage of olefins.

SUMMARY

According to one embodiment, the invention relates to a process for converting synthesis gas to a hydrocarbon mixture comprising contacting a feed comprising a mixture of carbon monoxide and hydrogen with, in sequence:

a synthesis gas conversion catalyst in an upstream bed, wherein a first intermediate hydrocarbon mixture containing olefins and C₂₁₊ normal paraffins is formed over the synthesis gas conversion catalyst,

a hydroisomerization catalyst containing a metal promoter and an acidic component in an intermediate catalyst bed downstream of the upstream catalyst bed, wherein said C₂₁₊ normal paraffins of the first intermediate hydrocarbon mixture are hydroisomerized over the hydroisomerization catalyst thus forming a second intermediate hydrocarbon mixture containing olefins and no greater than about 5 wt % C₂₁₊ normal paraffins, and

an olefin saturation catalyst in a downstream catalyst bed downstream of the intermediate catalyst bed, wherein said olefins are saturated over the olefin saturation catalyst, thereby resulting in a final hydrocarbon mixture containing no greater than about 25 wt % olefins and containing no greater than about 5 wt % C₂₁₊ normal paraffins.

According to another embodiment, the invention relates to a process for converting synthesis gas to a hydrocarbon mixture comprising contacting a feed comprising a mixture of carbon monoxide and hydrogen with, in sequence:

a synthesis gas conversion catalyst in an upstream bed, wherein an intermediate hydrocarbon mixture containing olefins and C₂₁₊ normal paraffins is formed over the synthesis gas conversion catalyst, and

a mixture of an olefin saturation catalyst and a hydroisomerization catalyst containing a metal promoter and an acidic component in a downstream catalyst bed, whereby said C₂₁₊ normal paraffins of the intermediate hydrocarbon mixture are hydroisomerized and said olefins are saturated thus forming a hydrocarbon mixture containing no greater than about 25 wt % olefins and no greater than about 5 wt % C₂₁₊ normal paraffins.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic diagram illustrating a process for converting synthesis gas to liquid hydrocarbons according to an embodiment of the invention.

FIG. 2 is a schematic diagram illustrating a process for converting synthesis gas to liquid hydrocarbons according to another embodiment of the invention.

FIG. 3 is a schematic diagram illustrating a process for converting synthesis gas to liquid hydrocarbons according to yet another embodiment of the invention.

DETAILED DESCRIPTION

Referring to FIG. 1, a process is disclosed for the synthesis of a liquid hydrocarbon product 30 in the distillate fuel and/or lube base oil range from a feed of synthesis gas 2. Within a fixed bed reactor, multiple, small-diameter tubes (not shown) are enclosed in a cooling medium, e.g., steam or water. Provided within the process is a method for synthesizing a mixture of olefinic and paraffinic hydrocarbons by contacting the synthesis gas with a synthesis gas conversion catalyst in a first, upstream catalyst bed 4 within the reactor. The hydrocarbon mixture so formed can range from methane to light wax, containing only trace amounts (<0.5 wt %) of carbon numbers above 30, and may include linear, branched and cyclic compounds. As defined herein, the terms “wax” and “solid wax” refer to C₂₁₊ normal paraffins. The terms “Fischer-Tropsch wax” and “C₂₁₊ wax” are also used herein interchangeably to refer to C₂₁₊ normal paraffins. The hydrocarbon mixture formed in catalyst bed 4 is then contacted within the same reactor with a second, intermediate catalyst bed 6. The intermediate bed can include a hydrogenation catalyst for hydrogenating olefins and a catalyst for hydroisomerizing the straight chain hydrocarbons. The upstream bed performs synthesis gas conversion while the intermediate bed performs hydroisomerization and optional hydrocracking. The synthesis gas conversion and the subsequent hydroisomerization can conveniently be carried out in a single reactor under essentially common reaction conditions without having to provide a separate reactor for hydroisomerization and optional hydrocracking. By “essentially common reaction conditions” is meant that the temperature of the cooling medium within the reactor 10 is constant from one point to another within a few degrees Celsius (e.g., 0-3° C.) and the pressure within the reactor is allowed to equilibrate between the two beds. Optionally, although not preferably, more than one cooling system may be used utilizing more than one cooling medium physically separated from each other, in which case the cooling media may be at differing temperatures. The temperatures and pressures of the upstream bed 4 and intermediate bed 6 can differ somewhat, although advantageously it is not necessary to separately control the temperature and pressure of the two beds. The bed temperatures will depend on the relative exotherms of the reactions proceeding within them. Exotherms generated by synthesis gas conversion are greater than those generated by hydrocracking; therefore in the case of constant reactor tube diameter, the average upstream bed temperature will generally be higher than the average intermediate bed temperature. The temperature difference between the beds will depend on various reactor design factors, including, but not limited to, the type and temperature of the cooling medium, the diameter of the tubes in the reactor, the rate of gas flow through the reactor, and so forth. For adequate thermal control, the temperatures of the two beds are preferably maintained within about 10° C. of the cooling medium temperature, and therefore the difference in temperature between the upstream and intermediate beds is preferably less than about 20° C., even less than about 10° C. The pressure at the end of the upstream bed is equal to the pressure at the beginning of the intermediate bed since the two beds are open to one another. Note that there will be a pressure drop from the top of the upstream bed to the bottom of the intermediate bed because gas is being forced through narrow tubes within the reactor. The pressure drop across the reactor could be as high as about 50 psi (3 atm), therefore the average difference in pressure between the beds could be up to about 25 psi (1.5 atm).

A feed of synthesis gas 2 is introduced to the reactor via an inlet (not shown). The ratio of hydrogen to carbon monoxide of the feed gas is generally high enough that productivity and carbon utilization are not negatively impacted by not adding hydrogen in addition to the hydrogen of the syngas into the reactor or producing additional hydrogen using water-gas shift. The ratio of hydrogen to carbon monoxide of the feed gas is also generally below a level at which excessive methane would be produced. Advantageously, the ratio of hydrogen to carbon monoxide is between about 1.0 and about 2.2, even between about 1.5 and about 2.2. If desired, pure synthesis gas can be employed or, alternatively, an inert diluent, such as nitrogen, CO₂, methane, steam or the like can be added. The phrase “inert diluent” indicates that the diluent is non-reactive under the reaction conditions or is a normal reaction product. It may be desirable to operate the syngas conversion process in a partial conversion mode, for instance 50-60 wt % based on CO.

According to one embodiment of the present process, the upstream bed 4 contains a Fischer-Tropsch synthesis gas conversion catalyst. The Fischer-Tropsch synthesis gas conversion catalyst can be cobalt, iron or ruthenium, or mixtures including cobalt, iron or ruthenium. Catalysts having low water gas shift activity and suitable for lower temperature reactions, such as cobalt and ruthenium, are preferred. The synthesis gas conversion catalyst can be supported on any suitable support, such as solid oxides, including but not limited to alumina, silica or titania or mixtures thereof. As nonlimiting examples, the synthesis gas conversion catalyst can be present on the support in an amount between 5% and 50% by weight in the case of cobalt, and between 0.01% and 1% by weight in the case of ruthenium.

According to another embodiment of the present process, the upstream bed 4 contains a hybrid synthesis gas conversion catalyst. As used herein, the phrases “hybrid Fischer-Tropsch catalyst” and “hybrid synthesis gas conversion catalyst” are used interchangeably to refer to a Fischer-Tropsch catalyst comprising a Fischer-Tropsch component as well as a component containing the appropriate functionality to convert in a single-stage the primary Fischer-Tropsch products into desired products (i.e., minimize the amount of heavier, undesirable products). A hybrid synthesis gas conversion catalyst contains a synthesis gas conversion catalyst in combination with an olefin isomerization catalyst, for example a relatively acidic zeolite, for isomerizing double bonds in C₄₊ olefins as they are formed. Methods for preparing a hybrid catalyst of this type are described in co-pending U.S. patent application Ser. No. 12/343,534, incorporated herein in its entirety by reference. Co-pending U.S. patent application Ser. No. 12/343,534 describes a method comprises impregnating a zeolite extrudate using a solution comprising a cobalt salt to provide an impregnated zeolite extrudate and activating the impregnated zeolite extrudate by a reduction-oxidation-reduction cycle. According to the present process, the Fischer-Tropsch component (also referred to as “FT component” or “FT metal”) is not necessarily cobalt, but may also comprise ruthenium, iron or mixtures including cobalt, iron or ruthenium. Impregnation of a zeolite using a substantially non-aqueous solution comprising an FT metal salt and a salt of a metal promoter, if desired, followed by activation by a reduction-oxidation-reduction cycle reduces ion-exchange with zeolite acid sites, thereby increasing the overall activity of the zeolite component. The resulting catalyst comprises FT metal distributed as small crystallites upon the zeolite support. The zeolite support, impregnation method and reduction-oxidation-reduction cycle used to activate the catalyst are described in detail below.

The use of zeolite extrudates as the zeolite support is beneficial, for the relatively larger zeolite extrudate particles result in lower pressure drop and are subject to less attrition than zeolite powder or even granular zeolite (e.g., having a particle size of about 300-1000 micrometers). Methods of formation of zeolite extrudates are readily known to those of ordinary skill in the art. Wide variations in macroporosity are possible with such extrudates. For the hybrid synthesis gas conversion catalyst, without wishing to be bound by theory, it is believed that as high a macroporosity as possible, consistent with high enough crush strength to enable operation in long reactor tubes, will be advantageous in minimizing diffusion constraints on activity and selectivity.

A zeolite support is a molecular sieve that contains silica in the tetrahedral framework positions. Examples include, but are not limited to, silica-only (silicates), silica-alumina (aluminosilicates), silica-boron (borosilicates), silica-germanium (germanosilicates), alumina-germanium, silica-gallium (gallosilicates) and silica-titania (titanosilicates), and mixtures thereof.

Molecular sieves, in turn, are crystalline materials that have regular passages (pores). If examined over several unit cells of the structure, the pores will form an axis based on the same units in the repeating crystalline structure. While the overall path of the pore will be aligned with the pore axis, within a unit cell, the pore may diverge from the axis, and it may expand in size (to form cages) or narrow. The axis of the pore is frequently parallel with one of the axes of the crystal. The narrowest position along a pore is the pore mouth. The pore size refers to the size of the pore mouth. The pore size is calculated by counting the number of tetrahedral positions that form the perimeter of the pore mouth. A pore that has 10 tetrahedral positions in its pore mouth is commonly called a 10-ring pore. Pores of relevance to catalysis in this application have pore sizes of 8 rings or greater. If a molecular sieve has only one type of relevant pore with an axis in the same orientation to the crystal structure, it is called 1-dimensional. Molecular sieves may have pores of different structures or may have pores with the same structure but oriented in more than one axis related to the crystal. In these cases, the dimensionality of the molecular sieve is determined by summing the number of relevant pores with the same structure but different axes with the number of relevant pores of different shape.

Exemplary zeolite supports of the hybrid synthesis gas conversion catalyst include, but are not limited to, those designated SSZ-13, SSZ-33, SSZ-46, SSZ-53, SSZ-55, SSZ-57, SSZ-58, SSZ-59, SSZ-64, ZSM-5, ZSM-11, ZSM-12, TS-1, MTT (e.g., SSZ-32, ZSM-23 and the like), H-Y, BEA (zeolite Beta), SSZ-60 and SSZ-70. These molecular sieves each contain silicon as the major tetrahedral element, have 8 to 12 ring pores, and are microporous molecular sieves, meaning having pore mouths of 20 rings or less.

The zeolite supports can have an external surface area of between about 100 m²/g and about 300 m²/g, for example, about 180 m²/g. Micropore volumes for 80% ZSM-5 are between about 90 and 112 μL/g, with lower volumes implying some occlusion or loss of micropore structure. BET surface area is a sum of external area and micropore area. The zeolite supports can further have porosity of between about 30 and 80%, total intrusion volume of between about 0.25 and 0.60 cc/g, and crush strength of between about 1.25 and 5 lb/mm. Si/Al ratio of the zeolite component can be between about 10 and 100.

Initially, the zeolite support can be treated by oxidative calcination at a temperature in the range of from about 450° to about 900° C., for example, from about 600° to about 750° C. to remove water and any organics from the zeolite support.

Meanwhile, a non-aqueous organic solvent solution of a FT component salt, and, if desired, aqueous or non-aqueous organic solvent solutions of metal promoter salts, for example, are prepared. Any suitable salt, such as nitrate, chloride, acetate or the like can be used. Aqueous solutions for the promoters can be used in very small amounts. As used herein, the phrase “substantially non-aqueous” refers to a solution that includes at least 95 volume % non-aqueous solution. In general, any metal salt which is soluble in the organic solvent and will not have a poisonous effect on the catalyst can be utilized. The non-aqueous organic solvent is a non-acidic liquid which is formed from moieties selected from the group consisting of carbon, oxygen, hydrogen and nitrogen, and possesses a relative volatility of at least 0.1. The phrase “relative volatility” refers to the ratio of the vapor pressure of the solvent to the vapor pressure of acetone, as reference, when measured at 25° C. Suitable solvents include, for example, ketones, such as acetone, butanone (methyl ethyl ketone); the lower alcohols, e.g., methanol, ethanol, propanol and the like; amides, such as dimethyl formamide; amines, such as butylamine; ethers, such as diethylether and tetrahydrofuran; hydrocarbons, such as pentane and hexane; and mixtures of the foregoing solvents. Suitable cobalt salts include, for example, cobalt nitrate, cobalt acetate, cobalt carbonyl, cobalt acetylacetonate, or the like. Likewise, any suitable ruthenium salt, such as ruthenium nitrate, chloride, acetate or the like can be used. In an embodiment, ruthenium acetylacetonate is used. In general, any metal salt which is soluble in the organic solvent and will not have a poisonous effect on the metal catalyst or on the acid sites of the zeolite can be utilized.

The calcined zeolite support is then impregnated in a dehydrated state with the substantially non-aqueous, organic solvent solution of the metal salts. Thus, the calcined zeolite support should not be unduly exposed to atmospheric humidity so as to become rehydrated. Any suitable impregnation technique can be employed including techniques well known to those skilled in the art so as to distend the catalytic metals in a uniform thin layer on the catalyst zeolite support. For example, the FT component and promoter can be deposited on the zeolite support material by the “incipient wetness” technique. Such technique is well known and requires that the volume of substantially non-aqueous solution be predetermined so as to provide the minimum volume which will just wet the entire surface of the zeolite support, with no excess liquid.

Alternatively, the excess solution technique can be utilized if desired. If the excess solution technique is utilized, then the excess solvent present, e.g., acetone, is merely removed by evaporation. Multiple impregnations are often needed to achieve the desired metal loading, with intervening drying and calcination treatments to disperse and decompose the metal salts. The FT component content can be varied from about 0.5 weight % to about 25 weight %.

A promoter metal may be included in the hybrid synthesis gas conversion catalyst if desired. For example, when the FT component is cobalt, suitable promoters include, for example, ruthenium, platinum, palladium, silver, gold, rhenium, manganese and copper. When the FT component is ruthenium, suitable promoters include, for example, rhenium, platinum, palladium, silver, gold, manganese and copper. As an example, for a catalyst containing about 10 weight % cobalt, the amount of ruthenium promoter can be from about 0.01 to about 0.50 weight %, for example, from about 0.05 to about 0.25 weight % based upon total catalyst weight. The amount of ruthenium would accordingly be proportionately higher or lower for higher or lower cobalt levels, respectively. A catalyst level of about 10 weight % has been found to best for 80 weight % ZSM-5 and 20 weight % alumina. The amount of cobalt can be increased as amount of alumina increases, up to about 20 weight % Co.

Next, the substantially non-aqueous solution and zeolite support are stirred while evaporating the solvent at a temperature of from about 25° to about 50° C. until “dryness”. The impregnated catalyst is slowly dried at a temperature of from about 110° to about 120° C. for a period of about 1 hour so as to spread the metals over the entire zeolite support. The drying step is conducted at a very slow rate in air.

The dried catalyst may be reduced directly in hydrogen or it may be calcined first. The dried catalyst is calcined by heating slowly in flowing air, for example 10 cc/gram/minute, to a temperature in the range of from about 200° to about 350° C., for example, from about 250° to about 300° C., that is sufficient to decompose the metal salts and fix the metals. The aforesaid drying and calcination steps can be done separately or can be combined. However, calcination should be conducted by using a slow heating rate of, for example, 0.5° to about 3° C. per minute or from about 0.5° to about 1° C. per minute and the catalyst should be held at the maximum temperature for a period of about 1 to about 20 hours, for example, for about 2 hours.

The foregoing impregnation steps are repeated with additional substantially non-aqueous solutions in order to obtain the desired metal loading. Metal promoters can be added with the FT component, but they may be added in other impregnation steps, separately or in combination, before, after or between impregnations of FT component.

After the last impregnation sequence, the loaded catalyst zeolite support is then subjected to the ROR activation treatment comprising the steps, in sequence, of (A) reduction in hydrogen, (B) oxidation in an oxygen-containing gas, and (C) reduction in hydrogen, the activation procedure being conducted at a temperature below 500° C., even below 450° C., even below 400° C., even below 300° C., depending on the FT component being used. Temperatures between 100° and 450° C., even between 250° and 400° C., are suitable for the reduction steps. The oxidation step is between 200° and 300° C. These activation steps are conducted while heating at a rate of from about 0.1° to about 5° C., for example, from about 0.1° to about 2° C. It has been found that the activation procedure provides a catalyst with improved reaction rates when the catalyst is prepared by impregnation of a zeolite support with an FT component such as cobalt or ruthenium. Moreover, the activation procedure can significantly improve the activity of the catalyst when a promoter has been previously added.

The ROR activation procedure of the present disclosure is now described in more detail. The impregnated catalyst can be slowly reduced in the presence of hydrogen. If the catalyst has been calcined after each impregnation, to decompose nitrates or other salts, then the reduction may be performed in one step, after an inert gas purge, with heating in a single temperature ramp (e.g., 1° C./min.) to the maximum temperature and held at that temperature, from about 250° or 300° to about 450° C., for example, from about 350° to about 400° C., for a hold time of 6 to about 65 hours, for example, from about 16 to about 24 hours. Pure hydrogen is preferred in the first reduction step. If nitrates are still present, the reduction can be conducted in two steps wherein the first reduction heating step is carried out at a slow heating rate of no more than about 5° C. per minute, for example, from about 0.1° to about 1° C. per minute up to a maximum hold temperature of 200° to about 300° C., for example, 200° to about 250° C., for a hold time from about 6 to about 24 hours, for example, from about 16 to about 24 hours under ambient pressure conditions. In the second treating step of the first reduction, the catalyst can be heated at from about 0.5° to about 3° C. per minute, for example, from about 0.1° to about 1° C. per minute to a maximum hold temperature of from about 250° or 300° up to about 450° C., for example, from about 350° to about 400° C. for a hold time of 6 to about 65 hours, for example, from about 16 to about 24 hours. Although pure hydrogen is preferred for these reduction steps, a mixture of hydrogen and nitrogen can be utilized.

The reduction may involve the use of a mixture of hydrogen and nitrogen at 100° C. for about one hour; increasing the temperature 0.5° C. per minute until a temperature of 200° C.; holding that temperature for approximately 30 minutes; and then increasing the temperature 1° C. per minute until a temperature of 350° C. is reached and then continuing the reduction for approximately 16 hours. Reduction should be conducted slowly enough and the flow of the reducing gas maintained high enough to maintain the partial pressure of water in the off-gas below 1%. Before and after all reductions, the catalyst is purged in an inert gas such as nitrogen, argon or helium.

The reduced catalyst is passivated at ambient temperature (25°-35° C.) by flowing diluted air over the catalyst slowly enough so that a controlled exotherm of no larger than +50° C. passes through the catalyst bed. After passivation, the catalyst is heated slowly in diluted air to a temperature of from about 300° to about 350° C. (preferably 300° C.) in the same manner as previously described in connection with calcination of the catalyst.

Next, the reoxidized catalyst is then slowly reduced again in the presence of hydrogen, in the same manner as previously described in connection with the initial reduction of the impregnated catalyst. This reduction may be accomplished in a single temperature ramp and held, as described above, for reduction of calcined catalysts.

While the ROR activation procedure of the present disclosure may be used to improve activity of the hybrid synthesis gas conversion catalyst of the present disclosure, any technique well known to those having ordinary skill in the art to distend the catalytic metals in a uniform manner on the catalyst zeolite support is suitable, provided they do not promote ion exchange with zeolite acid sites.

The hybrid synthesis gas conversion catalyst has an average particle diameter of from about 0.01 to about 6 millimeters; for example, from about 1 to about 6 millimeters.

According to yet another embodiment, the upstream bed 4 contains a mixture of conventional Fischer-Tropsch catalyst and a hybrid synthesis gas conversion catalyst, wherein the bed contains between about 1 and about 99 weight % conventional Fischer-Tropsch catalyst and about 1 and about 99 weight % hybrid synthesis gas conversion catalyst, based on total catalyst weight.

The intermediate catalyst bed 6 contains a hydroisomerization catalyst for hydroisomerizing straight chain hydrocarbons. The hydroisomerization catalyst is a bifunctional catalyst containing a hydrogenation component comprising a metal promoter and an acidic component. The hydroisomerization catalyst can be selected from 10-ring and larger zeolites. Suitable materials for use as the hydroisomerization catalyst include, as not limiting examples, SSZ-32, ZSM-57, ZSM-48, ZSM-22, ZSM-23, SAPO-11 and Theta-1. The hydroisomerization catalysts can also be non-zeolitic materials.

According to one embodiment, the intermediate catalyst bed 6 also contains a hydrocracking catalyst for cracking straight chain hydrocarbons. The hydrocracking catalyst is an acid catalyst material and can be a material such as amorphous silica-alumina or tungstated zirconia or a zeolitic or non-zeolitic crystalline molecular sieve. The crystallographic free diameters of the channels of molecular sieves are published in the “Atlas of Zeolite Framework Types”, Fifth Revised Edition, 2001, by C H. Baerlocher, W. M. Meier, and D. H. Olson. Elsevier, pp 10-15, which is incorporated herein by reference. Examples of suitable hydrocracking medium pore molecular sieves include zeolite Y, zeolite X and the so called ultra stable zeolite Y and high structural silica:alumina ratio zeolite Y such as for example described in U.S. Pat. Nos. 4,401,556, 4,820,402 and 5,059,567, herein incorporated by reference. The phrase “medium pore” as used herein means having a crystallographic free diameter in the range of from about 3.9 to about 7.1 Angstrom when the molecular sieve is in the calcined form. Small crystal size zeolite Y, such as described in U.S. Pat. No. 5,073,530, herein incorporated by reference, can also be used. Other zeolites which show utility as cracking catalysts include those designated as SSZ-13, SSZ-33, SSZ-46, SSZ-53, SSZ-55, SSZ-57, SSZ-58, SSZ-59, SSZ-64, ZSM-5, ZSM-11, ZSM-12, ZSM-23, H-Y, beta, mordenite, SSZ-74, ZSM-48, TON type zeolites, ferrierite, SSZ-60 and SSZ-70. Non-zeolitic molecular sieves which can be used include, for example, silicoaluminophosphates (SAPO) such as SAPO-11, SAPO-31 and SAPO-41, ferroaluminophosphate, titanium aluminophosphate and the various ELAPO molecular sieves described in U.S. Pat. No. 4,913,799 and the references cited therein. Details regarding the preparation of various non-zeolite molecular sieves can be found in U.S. Pat. No. 5,114,563 (SAPO); U.S. Pat. No. 4,913,799 and the various references cited in U.S. Pat. No. 4,913,799, hereby incorporated by reference in their entirety. Mesoporous molecular sieves can also be included, for example the M41S family of materials (J. Am. Chem. Soc. 1992, 114, 10834-10843), MCM-41 (U.S. Pat. Nos. 5,246, 689, 5,198,203, 5,334,368), and MCM48 (Kresge et al., Nature 359 (1992) 710).

The hydrocracking and hydroisomerization catalysts can optionally contain a metal promoter and a cracking component. The metal promoter is typically a metal or combination of metals selected from Group VIII noble and non-noble metals, Group 1B coinage metals, and Group VIB metals. Noble and coinage metals which can be used include platinum, palladium, rhodium, ruthenium, osmium, silver, gold and iridium, or any combination thereof. Non-noble metals which might be used include molybdenum, tungsten, nickel, cobalt, copper, rhenium, or any combination thereof.

The metal promoter can be incorporated into the catalyst mixture by any one of numerous procedures. It can be added either to the cracking component, to the support or a combination of both. In the alternative, the Group VIII components can be added to the cracking component or matrix component by co-mulling, impregnation, or ion exchange and the Group VI components, i.e., molybdenum and tungsten can be combined with the refractory oxide by impregnation, co-mulling or co-precipitation. These components are usually added as a metal salt which can be thermally converted to the corresponding oxide in an oxidizing atmosphere or reduced to the metal with hydrogen or other reducing agent.

According to one embodiment, the intermediate catalyst bed 6 contains a combination of a hydroisomerization component, e.g. a noble metal-promoted zeolite of the SSZ-32 family and a solid acid hydrocracking component, e.g. Pd/ZSM-5. The proportion of cracking and hydroisomerization catalysts in the intermediate bed is advantageously optimized to balance the isomerization activity with the cracking activity. If there is excessive cracking catalyst the resulting product may be lighter than desired. The cracking catalyst converts the n-paraffin wax product to a suitable chain length while the hydroisomerization component isomerizes the n-paraffin product, resulting in an entirely liquid isomerized product. If the desire is to produce a heavier, diesel range product, then the catalyst combination should exhibit less cracking and more isomerization. By including Pd/SSZ-32, for example, it has been found that more isomerization can be achieved. If there is insufficient cracking catalyst the hydroisomerization catalyst may be unable to convert the wax to liquid products under the mild process conditions of the present process. Accordingly, it may be advantageous to include in the intermediate bed a combination of both a cracking catalyst component and a hydroisomerization catalyst in the correct proportions so as to obtain a desired product, e.g. having an average molecular weight in the diesel range, i.e. C₁₁ to C₂₀, and containing no solid wax phase at ambient conditions.

The amount of catalyst mixture needed in the intermediate bed will in part depend on the tendency of the synthesis gas conversion catalyst in the upstream bed to produce wax and will in part depend on process conditions. In general, the weight of the catalyst mixture in the intermediate bed is between about 0.5 and about 2.5 times the weight of the catalyst in the upstream bed.

The reaction temperature is suitably from about 160° C. to about 260° C., for example, from about 175° C. to about 250° C. or from about 185° C. to about 235° C. Higher reaction temperatures favor lighter products. The total pressure is, for example, from about 1 to about 100 atmospheres, for example, from about 3 to about 35 atmospheres or from about 5 to about 20 atmospheres. Higher reaction pressures favor heavier products. The gaseous hourly space velocity based upon the total amount of feed is less than 20,000 volumes of gas per volume of catalyst per hour, for example, from about 100 to about 5000 v/v/hour or from about 1000 to about 2500 v/v/hour.

Fixed bed reactor systems have been developed for carrying out the Fischer-Tropsch reaction. Such reactors are suitable for use in the present process. For example, suitable Fischer-Tropsch reactor systems include multi-tubular fixed bed reactors the tubes of which are loaded with catalyst.

The process over the upstream and intermediate beds provides for a high yield of paraffinic hydrocarbons in the middle distillate and/or light base-oil range under essentially common reaction conditions. The intermediate mixture produced 3 is liquid at about 0° C. The mixture 3 is substantially free of solid wax by which is meant that the product is a single liquid phase at ambient conditions without the visibly cloudy presence of an insoluble solid wax phase. By “ambient conditions” is meant a temperature of 15° C. and a pressure of 1 atmosphere. The intermediate mixture 3 has the following composition:

0-20, for example, 5-15 or 8-12, weight % CH₄;

0-20, for example, 5-15 or 8-12, weight % C₂-C₄;

60-95, for example, 70-90 or 76-84, weight % C₅₊; and

0-5 weight % C₂₁₊ normal paraffins.

In a typical Fischer-Tropsch process, the product obtained is a predominantly a normal or linear paraffin product, meaning free of branching. If the C₂₁₊ fraction present within a predominantly linear product is greater than 5 weight %, the product has been found to contain a separate, visible solid wax phase. Products of the present process may actually contain C₂₁₊ at greater than 5 weight % without a visible solid wax phase. This is believed to be because of the hydroisomerization capability of the hydroisomerization catalyst. Branched paraffins have lower melting points compared with normal or linear paraffins such that products of the present process can contain a greater percentage of C₂₁₊ fraction and still remain a liquid which is free of a separate, visible solid wax phase at ambient conditions. The present process provides a product having a concentration of isomerized (i.e., containing at least single branches) C₂₁₊ paraffin of at least 30 weight % based on the weight of the C₂₁₊ fraction (as determined by gas chromatography). The result is a product which is liquid and pourable at ambient conditions. Liquid hydrocarbons produced by the present process advantageously have a cloud point as determined by ASTM D 2500-09 of 15° C. or less, even 10° C. or less, even 5° C. or less, and even as low as 2° C.

In addition, the present process provides for a high yield of paraffinic hydrocarbons in the middle distillate and/or light base-oil range without the need for separation of products arising from the first catalyst bed and without the need for a second reactor containing catalyst for hydrocracking and/or hydroisomerization. Process water arising from the first catalyst bed is not required to be separated from the reactor during the hydroisomerization of said C₂₁₊ normal paraffins. It has been found that with a proper combination of catalyst composition, catalyst bed placement and reaction conditions, both the synthesis gas conversion reaction and the subsequent hydrocracking and/or hydroisomerization reactions can be conducted within a single reactor under essentially common process conditions.

Intermediate mixture 3 is optionally directed to a separator 15 which utilizes a drop in temperature to condense water 17 and separate intermediate mixture 21 and gas stream 19. Gas stream 19 is recycled to the upstream bed via a compressor (not shown). A portion of gas stream 19 is optionally sent to a syngas generation unit, e.g. an autothermal reformer (not shown), or flared (not shown) in order to reduce the unconverted gas content recycled to the upstream bed.

Intermediate mixture 3 (or intermediate mixture 21 when optional separator 15 is used) is passed to downstream catalyst bed 16 containing an olefin saturation catalyst capable of saturating the olefins contained in the intermediate mixture. Hydrogen stream 28 is fed to catalyst bed 16. The resulting product stream 30 contains no greater than about 25 weight percent olefins, preferably no greater than about 5 weight percent olefins, even essentially no olefins. By “essentially no olefins” is meant that the product stream has a bromine number of less than about 1.

According to this embodiment, the downstream catalyst bed 16 can contain a catalyst comprised of a hydrogenation component useful as an olefin saturation catalyst deposited on a support. The hydrogenation component can be a Group IB noble metal, a Group VIII noble metal, or a combination thereof. Preferred noble metals include platinum, palladium, rhodium, iridium, silver, osmium and gold, and combinations thereof, and metal combinations including ruthenium. According to this embodiment, the downstream catalyst bed 16 can also include an olefin saturation catalyst having a hydrogenation component selected from a metal or combination of Group VIII non-noble metals and Group VIB metals. Non-noble metals which can be used include molybdenum, tungsten, nickel, iron, zinc, copper, lead and cobalt. Suitable combinations of metals include at least one Group VIII metal and one Group VIB metal, e.g., nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, and cobalt-tungsten. Preferred non-noble metal overall catalyst compositions contain in excess of about 5 weight percent, preferably about 5 to about 40 weight percent molybdenum and/or tungsten, and at least about 0.5, and generally about 1 to about 15 weight percent of nickel and/or cobalt determined as the corresponding oxides. The non-noble metal hydrogenation metals can be present in the olefin saturation catalyst composition as metal sulfides when such compounds are readily formed from the particular metal involved. The sulfide form of these metals may have desirable activity, selectivity and activity retention.

The olefin saturation catalyst can be supported on any suitable support, such as solid oxides, including but not limited to alumina, silica or titania, or mixtures thereof. This support may be a zeolite support containing silica in the tetrahedral framework positions. Examples include, but are not limited to, silica only (silicates) such as silicalite, silica-alumina (aluminosilicates), silica-boron (borosilicates), silica-germanium (germanosilicates), aluminum-germanium, silica-gallium (gallosilicates) and silica-titania (titanosilicates), and mixtures thereof.

According to this embodiment, the pressure in the downstream catalyst bed 16 is between about 200 psig (1.4 MPa) and about 3000 psig (21 MPa), preferably between about 500 psig (3.4 MPa) and about 2000 psig (13 MPa). Temperature ranges in the downstream catalyst bed 16 are usually between about 300° F. (150° C.) and about 700° F. (370° C.), preferably between about 400° F. (205° C.) and about 500° F. (260° C.). The LHSV is usually within the range of from about 0.2 to about 2.0 h⁻¹, preferably 0.2 to 1.5 h⁻¹andmost preferably from about 0.7 to 1.0 h⁻¹. Hydrogen is usually supplied to the downstream catalyst bed 16 at a rate of from about 1000 SCF (28 m³) to about 10,000 SCF (280 m³) per barrel of feed. Typically the hydrogen is fed at a rate of about 3000 SCF (85 m³) per barrel of feed.

According to another embodiment, the upstream, intermediate and downstream beds are located in series physically within the same reactor. FIG. 2 illustrates this embodiment, in which synthesis gas feed 2 enters the reactor and passes over upstream catalyst bed 4 of synthesis gas conversion catalyst, intermediate catalyst bed 6 of hydroisomerization catalyst, and downstream catalyst bed 16 of olefin saturation catalyst, each of the synthesis gas conversion catalyst, hydroisomerization catalyst and olefin saturation catalyst having been previously described herein. Stream 18 is optionally passed to separator 14 and stream 12 is optionally recycled.

According to yet another embodiment, the hydroisomerization catalyst and olefin saturation catalyst are physically mixed within the same bed 10, within the same reactor downstream of catalyst bed 4. FIG. 3 illustrates this embodiment, in which synthesis gas feed 2 enters the reactor and passes over upstream catalyst bed 4 of synthesis gas conversion catalyst, and downstream catalyst bed 10 of hydroisomerization catalyst and olefin saturation catalyst, as previously described herein. Downstream bed 10 can contain between about 10 and about 90 weight % hydroisomerization catalyst and between about 90 and about 10 weight % olefin saturation catalyst, respectively, based on total catalyst weight within bed 10. Stream 18 is optionally passed to separator 14 and stream 12 is optionally recycled.

EXAMPLES

In the following examples and comparative examples, a 5 mm ID, 300 mm long stainless steel tube reactor was used. The reactor included a 140 mm long heated zone. Catalyst was held in the central 80-100 mm of the heated zone.

The catalysts were subjected to the following ROR activation procedure. The initial reduction in 96% H2-4% Ar at 10 atm was conducted by ramping from ambient temperature to 350° C. at 0.5° C./minute, with intermediate temperature holds at 150° C. and 250° C. for 3 hours each, and the reduction was held at 350° C. for 12 hours. The flow was then switched to 96% N2-4% Ar for purging and the temperature was dropped to 100° C. over 12-15 hours and held at 100° C. for another 5 hours. Oxidation was then conducted by switching to a flow of 5% O2-91% N2-4% Ar and heating at 1° C./minute to 300° C., then holding at that temperature for another 6 hours. The flow was returned to 96% N2-4% Ar and the temperature was decreased to 100° C. over 10-12 hours. The catalysts were held at that temperature for an additional 16 hours. The final reduction was conducted in 96% H2-4% Ar, at 10 atm and 62.5 sccm flow. The heating ramp was 0.25° C./minute for this step. The final hold temperature was 250° C. After holding at 250° C. for 12 hours, the temperature was decreased to about 150° C. for the start of the test run.

The testing was started at 150° C. in a flow of 64% H2-32%CO-4% Ar (H2/CO=2). The pressure was 10 atm and the flow was 2.4 NL/h per position. The temperature was increased to 180° C. over the first several hours of operation. Thereafter, temperatures were never increased any faster than 0.01° C./min. when they were being raised. Temperatures were varied from 180° C. to 230° C., the H2/CO ratio from 1.5 to 2.0, the pressure from 10 to 20 atm and the flow rate from 1.2 to 2.4 NL/reactor/h. H2, CO, CO2, and total hydrocarbons were analyzed on line by ABB analyzers. C1-C12 hydrocarbons were analyzed on line by gas chromatography (FID- and TCD based); full isomer distributions were possible through C5. C13+ hydrocarbons were collected in hot traps and later analyzed offline in high-temperature GC equipment designed for wax analysis. Products out to C65 were analyzed if they were present. Bromine number was determined by test ASTM D 1159.

Example 1

150 mg of a syngas conversion (FT) catalyst comprising 7.5 weight % Co and 0.19 weight % Ru supported on 80 weight % ZSM-12 and 20 weight % alumina was prepared as follows, sized to 125-160 pm, and diluted four times by volume with carborundum. The FT catalyst was placed in the reactor, upstream of a mixture of 237 mg Pt/ZSM-5 and 253 mg Pd/SSZ-32, prepared as follows. The total length of both beds was 75 mm. The corresponding volume was 1.47 mL.

Preparation of FT Catalyst

A three-step incipient wetness impregnation method was used to prepare the catalyst. A solution was prepared by dissolving 125.824 g of cobalt(II) nitrate hexahydrate (obtained from Sigma-Aldrich), 2.041 g of ruthenium(III) nitrosyl nitrate (obtained from Alfa Aesar) and 3.381 g lanthanum (III) nitrate hexahydrate (obtained from Sigma-Aldrich) in water. 100 g of Puralox™ alumina SBA 200 (obtained from Sasol) support, after calcination in air at 750° C. for 2 h, was impregnated by using one-third of this solution to achieve incipient wetness. The prepared catalyst was then dried in air at 120° C. for 16 hours in a box furnace and then it was subsequently calcined in air by raising its temperature at a heating rate of 1° C./min to 300° C. and holding it at that temperature for 2 hours before cooling it back to ambient temperature. The above procedure was repeated to obtain the following loading of Co, Ru and La₂O₃ on the support: 20 wt % Co, 0.5% Ru and 1 wt % La₂O₃ and 78.5% alumina.

Preparation of Pt/ZSM-5 Catalyst

1.2466 g of tetraammineplatimum(II) nitrate (obtained from Aldrich) was dissolved in 60 cc of water. The resulting solution was added to 125.1 g CBV 8014 zeolite (obtained from Zeolyst International) extrudates. Most of the water was removed in a rotary evaporator under vacuum by heating slowly to 65° C. The vacuum-dried material was then further dried in an oven at 120° C. overnight. The dried catalyst was calcined at 300° C. for 2 hours in a muffle furnace.

Preparation of Pd/SSZ-32 Catalyst

28.34 g of tetraamminepalladium(II) nitrate solution, 10 wt % in H₂O (obtained from Aldrich) was dissolved in 100 cc of water. The resulting solution was added to 100 g of SSZ-32 base zeolite available from Chevron USA Inc. Most of the water was removed in a rotary evaporator under vacuum by heating slowly to 65° C. The vacuum-dried material was then further dried in an oven at 120° C. overnight. The dried catalyst was calcined at 300° C. for 2 hours in a muffle furnace.

Comparative Example 1

150 mg of the CoRu/alumina catalyst was diluted four times by volume as described in Example 1 and placed in the upstream bed. The total length of catalyst bed was 70 mm. The corresponding volume was 1.37 mL.

Comparative Example 2

150 mg of the CoRu/alumina catalyst was diluted four times by volume as described in Example 1 and placed in the upstream bed. 473 mg of the Pt/ZSM-5 catalyst described in Example 1 was placed in the downstream bed. The total length of the upstream and downstream catalyst beds was 70 mm. The corresponding volume was 1.37 mL.

The results are given in Table 1, including three runs for each example and comparative example. The results demonstrate the efficacy of the present process to significantly reduce olefin content of an intermediate hydrocarbon mixture.

Ex. 1 Ex. 1 Ex. 1 Comp. Ex. 1 Comp. Ex. 1 Comp. Ex. 1 Comp. Ex. 2 Comp. Ex. 2 Comp. Ex. 2 Conditions TOS, h 360 850 950 360 850 950 360 850 950 Temp, ° C. 220 220 220 220 220 220 220 220 220 Pressure, atm 10 10 20 10 10 20 10 10 20 Feed H₂/CO ratio 2.0 1.5 1.5 2.0 1.5 1.5 2.0 1.5 1.5 GHSV, SL/h/g 16 16 16 16 16 16 16 16 16 Results CO 30.8 17.2 15.3 28.7 16.4 13.5 29.7 16.6 14.7 conversion, % Rate, 4.36 2.92 3.40 4.18 2.81 3.27 4.27 2.85 3.34 gC/gCo/h Rate, 0.87 0.58 0.68 0.84 0.56 0.65 0.85 0.57 0.67 gC/gFTS/h Products (wt %) CH₄ 12% 11% 15% 12% 10% 15% 13% 11% 16% C₂  2%  2%  2%  2%  1%  2%  2%  2%  2% C₃  5%  5%  6%  4%  4%  6%  4%  4%  6% C₄-C₂₀ 76% 81% 63% 70% 70% 58% 77% 80% 61% C₂₁₊  5%  2% 14% 12% 15% 19%  5%  3% 15% % paraffin in C₅ fractions n-C₅ 98% 93% 100%  47% 32% 33% 75% 65% 50% Iso-C₅ 70% 34% 41% 43% 56% na  4%  3%  0% Total C₅ 93% 71% 83% 47% 33% 33% 43% 29% 31% % Iso compounds Total C₅ 19% 37% 28%  3%  2%  0% 47% 58% 37% % Olefins in  7% 29% 17% 53% 67% 67 57% 71% 69% C₅ 

1. A process for converting synthesis gas to a hydrocarbon mixture comprising contacting a feed comprising a mixture of carbon monoxide and hydrogen with, in sequence: a. a synthesis gas conversion catalyst in an upstream bed, wherein a first intermediate hydrocarbon mixture containing olefins and C₂₁₊ normal paraffins is formed over the synthesis gas conversion catalyst, b. a hydroisomerization catalyst containing a metal promoter and an acidic component in an intermediate catalyst bed downstream of the upstream catalyst bed, wherein said C₂₁₊ normal paraffins of the first intermediate hydrocarbon mixture are hydroisomerized over the hydroisomerization catalyst thus forming a second intermediate hydrocarbon mixture containing olefins and no greater than about 5 wt % C₂₁₊ normal paraffins, and c. an olefin saturation catalyst in a downstream catalyst bed downstream of the intermediate catalyst bed, wherein said olefins are saturated over the olefin saturation catalyst, thereby resulting in a final hydrocarbon mixture containing no greater than about 25 wt % olefins and containing no greater than about 5 wt % C₂₁₊ normal paraffins.
 2. The process of claim 1 wherein the final hydrocarbon mixture contains no greater than about 5 wt % olefins.
 3. The process of claim 1 wherein the final hydrocarbon mixture contains essentially no olefins.
 4. The process of claim 1 wherein the upstream bed and the intermediate beds are within a single reactor and have an essentially common reactor temperature and an essentially common reactor pressure.
 5. The process of claim 4 wherein the single reactor further comprises the downstream bed.
 6. The process of claim 1 wherein the synthesis gas conversion catalyst comprises cobalt, iron or ruthenium on a solid oxide support.
 7. The process of claim 6 wherein the solid oxide support is selected from the group consisting of alumina, silica, titania and mixtures thereof.
 8. The process of claim 1 wherein the hydroisomerization catalyst comprises a zeolite of the SSZ-32 family.
 9. The process of claim 1 wherein the intermediate bed further comprises a hydrocracking catalyst selected from the group consisting of amorphous silica-alumina, tungstated zirconia, zeolitic molecular sieve and non-zeolitic crystalline molecular sieve.
 10. The process of claim 1 wherein the hydroisomerization catalyst further comprises a metal promoter selected from the group consisting of cobalt, nickel, copper, ruthenium, rhodium, rhenium, palladium, silver, osmium, iridium, platinum, gold, molybdenum, tungsten, and oxides, and combinations thereof.
 11. The process of claim 11 wherein the hydrocracking catalyst further comprises a metal promoter selected from the group consisting of cobalt, nickel, copper, ruthenium, rhodium, rhenium, palladium, silver, osmium, iridium, platinum, gold, molybdenum, tungsten, and oxides, and combinations thereof.
 12. The process of claim 1 wherein the synthesis gas conversion catalyst further comprises a promoter selected from the group consisting of ruthenium, rhenium, platinum, palladium, gold, and silver.
 13. The process of claim 1 wherein the gaseous hourly space velocity is between about 100 and about 5000 volumes of gas per volume of catalyst per hour.
 14. The process of claim 4 wherein the reactor pressure is between about 3 atmospheres and about 35 atmospheres.
 15. The process of claim 4 wherein process water is not separated from the reactor during the hydroisomerization of said C₂₁₊ normal paraffins.
 16. The process of claim 4 wherein no hydrogen in addition to the mixture of carbon monoxide and hydrogen is added to the reactor.
 17. The process of claim 1 wherein the hydrocarbon mixture is substantially free of solid wax at ambient conditions.
 18. The process of claim 1 wherein the hydrocarbon mixture has an isomerized C₂₁₊ paraffin concentration of at least 30 weight % based on the weight of the C₂₁₊ fraction.
 19. The process of claim 1 wherein the hydrocarbon mixture has a cloud point no greater than 15° C.
 20. The process of claim 1 wherein between the intermediate catalyst bed and the downstream catalyst bed, the second intermediate hydrocarbon mixture is passed through a separator and separated into gas which is recycled to the upstream catalyst bed, water which is removed and liquid hydrocarbons which are passed to the downstream catalyst bed.
 21. The process of claim 4 wherein the single reactor is a multi-tubular fixed bed reactor.
 22. The process of claim 1 wherein the olefin saturation catalyst is selected from the group consisting of metals selected from Group IB noble metals and Group VIII noble metals and combinations thereof.
 23. The process of claim 1 wherein the olefin saturation catalyst is selected from the group consisting of platinum, palladium, rhodium, iridium, silver, osmium and gold, and combinations thereof.
 24. The process of claim 20 wherein the olefin saturation catalyst is selected from the group consisting of metals selected from Group VIII noble and non-noble metals and Group VIB metals, and combinations thereof.
 25. The process of claim 20 wherein the olefin saturation catalyst is selected from the group consisting of molybdenum, tungsten, nickel, iron, zinc, copper, lead, cobalt, nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, and cobalt-tungsten.
 26. The process of claim 20 wherein the olefin saturation catalyst comprises a metal sulfide.
 27. The process of claim 1 wherein the olefin saturation catalyst comprises a refractory inorganic oxide support.
 28. The process of claim 27 wherein the refractory inorganic oxide support comprises a zeolite.
 29. The process of claim 1 wherein the olefin saturation catalyst comprises a zeolite comprising SSZ-32.
 30. The process of claim 1 wherein the upstream catalyst bed temperature is between about 160° C. and about 260° C.
 31. The process of claim 1 wherein the upstream catalyst bed temperature is between about 175° C. and about 250° C.
 32. The process of claim 1 wherein the upstream catalyst bed temperature is between about 185° C. and about 235° C.
 33. The process of claim 4 wherein the temperature of the upstream catalyst bed and the temperature of the intermediate catalyst bed differ by no more than about 20° C.
 34. The process of claim 1 wherein the final hydrocarbon mixture produced comprises: 0-20 wt % CH₄; 0-20 wt % C₂-C₄; and 60-95 wt % C₅₊.
 35. A process for converting synthesis gas to a hydrocarbon mixture comprising contacting a feed comprising a mixture of carbon monoxide and hydrogen with, in sequence: a. a synthesis gas conversion catalyst in an upstream bed, wherein an intermediate hydrocarbon mixture containing olefins and C₂₁₊ normal paraffins is formed over the synthesis gas conversion catalyst, and b. a mixture of an olefin saturation catalyst and a hydroisomerization catalyst containing a metal promoter and an acidic component in a downstream catalyst bed, whereby said C₂₁₊ normal paraffins of the intermediate hydrocarbon mixture are hydroisomerized and said olefins are saturated thus forming a hydrocarbon mixture containing no greater than about 25 wt % olefins and no greater than about 5 wt % C₂₁₊ normal paraffins. 